Process for isomerizing paraffin hydrocarbons using an aluminum halide-hydrocarbon complex catalyst



2,375,321 ING c. w. NYsEwANDER l-:TAL

May v8, 1945. l l PROCESS FOR-ISOMERIZING PARAFFIN HYDROCARBONS US COMPLEX CATALYST AN ALUMINUM HALIDE-HYDROCARBON Filed Dec nh wuwNwN,

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raocnss Fon rsoiunmzmo r nnocAitBoNs USING aN en A e a-HYDROCARBON conan miv ATA- Cecii W. Nysewander, Highland, and Nathan Fragen, Hammond, Ind., assignors to Standard @il ll'lompariy, indiana Y Chicago, lll., a corporation of .applicaties neeeeber i5, i941, sei-isi No. 422,978 d Claims. (Cl. 26o-'6835) This invention relates -to the catalytic conversion of hydrocarbons and pertains more particularly to an improved process for the conversion of normally straight-chain hydrocarbons to branched-chain` hydrocarbons involving the use of an active aluminum halide-hydrocarbon complex as a catalyst and an activator therefor.,

It is well-known that branched-chain Vparafnic hydrocarbons and naphthas containing them in substantial proportions are-very valuable as motor fuels, and particularly as airplane fuels, because of-their high antiknock values, better lead response, freedom from gum-forming tendencies, and high heat content per unit weight'of fuel. `It has been proposedto Droduce'such naphthas from substantially saturated liquid fractions which are rich in straight-chain paraiin hydrocarbons by the use of aluminum chloride or other active aluminum halide catalysts in the presence of an activator such as hydrogen chloride. In carrying out the conversion of straight-chain pected properties of long catalyst life with comparatively easy regeneration.

A The preferred complex is preparedby contacting a. normallyl liquid hydrocarbon fraction rich in paraiilnic hydrocarbonswith anhydrous-aluminum chloride or aluminum bromide in the presence of an' activator, such as a hydrogen halide or a compound yielding ahi/drogen halide under the reaction conditions at a temperapressures or slightly. above.

turel within the range from about 50 to about 225 F., preferably at a temperature in the range from about 100 Fb to about 150 F. and at atmospheric The hydrocarbon feed stock from which the catalyst is prepared is tially free of olenic hydrocarbons since these y paraiiln hydrocarbons to branched-chain paramn t hydrocarbons on a reiinery scale, 'it is extremely advantageous that the process be continuous, and

this fact makes it very desirable to have the catalyst in liquid form so that it can be. pumped readily through pipes, tubes and other apparatus, or be employed as a deep pool in vertical towers through which the hydrocarbons to be converted are passed. It has long been known that aluminum chloride in the presence of a. hydrocarbon such bas an olenic or aromatic hydrocarbon, under conversion' conditions, will gradually and sometimes very rapidly be converted into analumnum chloride-hydrocarbon complex which is a liquid and retains e portion er the activity ef the ried out moet eect'ively in the presence of an active aluminum hande-hydrocarbon complex produced by the action of an aluminum halide such as anhydrous aluminum chlorideh on a'normally liquid substantially saturated hydrocarbon fraction in the presence of an activator such as hydrogenchloride at a Arelatively lowtemperapreferably free of unsaturates, although up to about 2% by Weight of'aromatics, or even 5% by weight can be present without too great a deleterous edect. The fraction should also be sub-stantoo react to form an undesirable complex with the aluminum halide. Cycloparafdns and naphthenes can be included "in the hydrocarbon feed stock. We have found that although any hydrocarbon fraction rich in paraiilnic hydrocarbons is suitable for our purpose, those containing branched-chain paraihnic hydrocarbons form s. particularly desirable complex with the aluminum halide. We also v 4canbe used for our purpose. Generally speaiding, when hydrocarbons oi low molecular weight are' employed for the formation of an aluminum haiide-hydrocarbon complex, it is necessary to employ higher temperatures than those herefore set forth.V A particularly suitable feed stock for the production ,of our catalyst can be a ht naphtha obtained 'from the fractionation" of straight-run sasolines lfrom crude oils, particularly those of Continent, Pennsylvania. or Michigen origin. or from the tural sesoline obtained from naturalgas wells or from distillate type wells. Licht naphthas havinga distillation point of not more than about 180 F., andprei'- erably those having a 95% distillation point of note ture, and that such complexes exhibit the unex- 5s more than F., are particularly suitable. Another excellent source of hydrocarbon: fractions for1 our purposeis the so-called commercial isooctane prepared by such processes as vthe poly-- erizatlon of nominallyA` gaseous oleflns to the ersor trimers, or the corresponding coor interpolymer's,'these polymers having been hydrogenated to form saturated branched-chain hyif Weight percent .Aluminum 12.5

.Chlorine 44 .Carbon plus hydrogen. 43.5

Cl/al (atomic ratio=2.7)

From the ratio of chlorine to aluminum it Will. be seen that the chlorine atoms have not been replaced by hydrocarbon radicals, since otherwise the ratio of chlorine to aluminum would have been 'of the order of 2':1,to 1:1. An attempted analysis ofthe hydrocarbons separated from the complex by hydrolysis was not too revealing as to their structure but did indicate that the hydrojcarbons presentwere not joined to the aluminum halide in parafflnic form, i.v e., that the pomplex was not a loose combination of the aluminum' halide and the original hydrocarbon in its original plex can be formed in situ under the reaction conditions by the introduction of the aluminum halides either in the form of a, slurry or a solution. A portion of the feed stock is usually Aemoptimum yield of product of maximum octane number. Another object of our invention is to provide an improved process whereby a hydrocarbon fraction rich in straight-chain paraiiinic hydrocarbons is contacted with an aluminum i chloride-hydrocarbon complex'of the above deform. The 'particular analysis of the complex is not important for our purpose, however, except as it serves to distinguish thesecatalysts from e aluminum halide-hydrocarbon complexes formed by the interaction of aluminum halide'with, for example, olens or aromatics. Apparently parafflns, oleiins and aromatics each form a distinct type of hydrocarbon complex with aluminum halide, these complexes differing materially in their catalytic effect as wellas in physical and other chemical characteristics, and we wish specically to'exclude from use 'in our process the catalysts formed by the interaction oan aluminum chloride with olens or aromatics.

A particularly eifectivelatalyst has been found to be obtained when the paraii'inic hydrocarbon with which the aluminum halide is reacted to form the catalyst contains atleast two side. chains per molecule, i. e.; when highly branched isomericparailnic hydrocarbons are employed as the hydrocarbon feed stock in the catalyst preparation step. ,A catalyst of this nature has been prepared by stirring together at atmospheric.

pressure aquantity of anhydrous aluminum chloride with an excess of commercial isooctane at 120 F. to 140 F.' until a liquid complex resulted.

Gaseous hydrogen chloride was bubbled through A.

the reaction mixture and a reflux arrangementwas employed. During the `complex formation large amounts of isobutane were produced and the unconverted. hydrocarbon liquid contained scription under conditions of temperature and .tacted with Aan aluminum halide-hydrocarbon complex in two or more zones of varying reactio conditions, whereby -a fraction rich in.

branched-chain paranic hydrocarbons of high octane number is produced rapidly 'with maximum yield. A further object is to provide animproved process wherein a light parafllnic naphtha is subjected to the action of an aluminum halidety-pe catalyst in a plurality of reaction zones maintained under-different .conditions for the production of the maximum octane number in the most eiilcient manner. A still further object of our invention isto correlate operating conditions such as'tmperature, pressure, catalyst activity, etc. throughput a multistage isomeriza tion system, in such a manner as 'to produce maximum yields of the most desirable products with a. consumptionfof catalyst, activator and hydrogen, and lwith an extremely edicient arrangement of reaction zones. Additional objects and advantagesof our invention will become apparentas the description thereof proceeds, particularly when read in conjunction with the accompanying drawing which represents a schematic now diagramof apparatus suitable for carrying out one embodiment of our invention.

The feed stock for our isomerization process can be any substantially saturated naphtha rich v in straight-chain paraiiinic hydrocarbons; for exabout 50% of material boiling higher than the end point of the original isooctane. The complex itself was separated from the unreacted. hydrocarbons and from the unreacted aluminum chlo- Y tively :free of oleiinic and/or aromatic hydrocarbons and particularly if itis a light parafnnic naphtha containing hydrocarbons having less than seven carbon atoms per molecule, the comafmple it can. be a relatively 4pure normally liquid straight-chain or slightly branched-chain parain hydrocarbon, such as normal pentane, norpredominantly straight-run naphthas such as" those from MichiganPennsylvana and Mid- Continuent crude oils are preferable, since they are much more readily available. Another excellent feed stock is a highly paramnic naphtha The catalystc complex had a viscosity line fractions and so-called distillates are also produced by the Fischer-Tropsch process from carbon monoxide and hydrogen. Natural gasosuitable andare plentiful and inexpensive in some rproduction area's. Lt is very important that the feed stock be free or substantially free from aromatic hydrocarbons since they have been found to reduce the activity of the catalyst to a very marked degree and consequently to limit seriously the amount of conversion obtained per unit weight of catalyst. Our preferred feed stock therefore contains less than about 5% and prefy traction step or other treatment is necessary or desirable in order to reduce the aromatic content of the feed. Olefinic hydrocarbons are also undesirable and should not 'be present in more than very small amounts, while cycloparaiiins or naphthenic hydrocarbons can be tolerated in considerable quantities. It salso frequently desirable to limit or eliminate arainic hydrocarbons higher boiling than the hexanes. Nevertheless, the feed stocks should preferably contain at least 50% of paran hydrocarbons and those containing at least 80% of paraffin hydrocarbons are especially desirable.

In general, the naphtha feed stock should have a boiling range within the range from about 100 F. to about 500 F., although naphthahaving an initial boiling point as low as about 30 F. including about 25 to 30% by weight of butanes can be used. A particularly suitable feed is one prepared by the distillation and fractionation of a straight-run or natural gasoline stock to produce a light naphtha having an initial boiling point in the range from about 30 F. to about 90 F., or higher, and a 95% boiling point in the range from about 145 F. to about 158 F., preferably about 152 F., although fractions having a, 95%

Ahas been refilled it is brought back onstream in the same marinier as reactor C just described. j If the time required by the cleaning and relling.

operation plus the induction period requires too great a time, avfourth reactor or more can be used.

The reaction is carried out in each reactor in the presence of an activator, such as hydrogen chloride or a substance aording hydrogen chloride under the reaction conditions. Preferably the hydrogen halide is absorbed in the charging stock, and since only a .minor amount of the hydrogen chloride in the system is consumed, a large part of the activator will be obtained from recycle gases. The amount of hydrogen chloride present in each reactor is desirably in the range from about to about 30 pounds per barrel of charging stock to the reactor, usually within the point of about 180 F. or less, preferably not above 170 F., can be tolerated providing they are substantialiy free of aromatics.

Briefly stated, our process provides for thel treatment oi a feed stock of the above described characteristics in two or more reactors in series, these reactors being maintained at varioustemperatures depending upon the condition' of the catalyst therein. We have found that the maximum optimum improvement can be obtained at a temperature Within the range, for example, of 200 to 300 F. The rate of octane improvement, howevenlor the reaction rate is substantially faster at higher temperatures, as for example, from 300 to 400 F. Furthermore, the fresh catalyst requires a period of' timean induction periodin which to reach its peak activity, and this induction period may be shortened by an initial period of operation at the'higher temperature named above. Thus the general plan of operation will be as follows: The'naphtha passes first through reactor A which contains the partially C spent catalyst heated to the more elevated temapproximate range of 8 or 10 pounds per barrel, but most of this is recovered andreintroduced into the system.

The reaction is also preferably carried out in the presence of hydrogen, and under superatmospheric pressures within the approximate range of 100 to 3000 pounds per square inch, preferably about 500 to about 1500 pounds per square inch, a large part of the pressure being due to the partial pressure of' the hydrogen present. The amount of hydrogen required will vary somewhat with the temperature, pressure and hydrogen chloride concentration in the reaction zone, ranging from about 20 cubic feet per barrel of stock charged at low temperatures, low pressures and low hydrogen chloride concentrations to 200 or more cubic` feet per barrel at high temperatures, pressures and. hydrogen `chloride concentrations. Usually the actual hydrogen consumption ,will be about 100 cubic feet per barrel, but in order to insure the presence of the requisite amount of hydrogen in the reactor We prefer /tintroduce about 100 to 300, preferably about 200 cubic feet of hydrogen per barrel of charging stock. t

The make-up aluminum chloride which'is preferably in the form of a complex but nevertheless pietely spent. During these last stages of deacf is in the highervrange of temperature. Thusat this stage the reactor A is removed from the system for cleaning and refining with fresh cat-l alyst, reactor B is atthe elevated temperature and accomplishing a large .part of the octane improvement, while reactor C at the lower temperature-is completing the reaction to give the highestpossible octane improvement. When reactor yof complex present in the reactor, the space velocity should be within the approximate range of,

0.2 to a0 volumes of liquid feed per hour per vol-l ume of complex in the reactor, preferably about one-half volume of stock per hour per volume of catalyst.

.Referring now to the drawingfeeq stoer, which for purposesof this description may be a light paaihnic naphtha cong substantial quantities of Cs `and Ce hydrocarbons, is added through line i0, pump ii and line i0 to absorber i3 while activator 'is added through line i@ 'and pump l0 and fiows ctercurrent to the charging stock stream. .as activator we can er'nploy a hydrogen halide, such as hydrogenchloride or hydrogen bromide, or we can use -a substance aording a hydrogen halide under the reaction conditions. Included in such substances are the alkyl halides 'such as propyl halide, butyl halide, ethyl halide,

particularlythe chlorides, aswell as chloroform and si crsanic halides.- We can also employ naphtha having dissolved therein the hydrogen ated or .fresh catalyst from line 2 8. The catachlorine which, with the hydrocarbons present in the feed stock, will form hydrogen chloride under the reaction conditions, and we can also use water which will react with the catalyst present to yield hydrogen halide, although this latter expedient 5 not too desirable, since it deteriorates the catalyst to a very considerable extent. We can also -naphtha at various temperatures and pressures is shown in the following table:

vTable Weight per cent HCI in ight napbtha HCI pressure Absorber I3 is preferablymaintaned at a tem- 35 perature within the range from about 80 F. to

.about 150 F. and at a pressure within the range from about 100 pounds per square inch to about 300 pounds per square inch, for example about 250 pounds per square inch. A very considerable partei the hydrogen halide is obtained by absorbing ,the hydrogen chloride from the. gases from th system'iritroduced through line I6. The

chloride or'other activator is withdrawn through line I1 to line I8. Instead of using allof the charging stock forA absorbing the hydrogen chloride, a 'maior partl can be directed through line I2 to absorber I3, and the remainder 'directed to line I8 via lineri9, valves 20 and 2| in lines 'I2 and I9 respectively being adjusted for that pur-,- pose.

.We preferthatlthe isomerization reaction be carried out under a` considerable hydrogen pressure and for'this reason .hydrogen is added `from line 22 through valve 23 and compressor 24 to line I8.` The reaction is'preferably carried out at hydrogen pressures withinthe range from about 100 pounds per squareinch to 3000.pounds per square inch, preferably between 500 pounds per square inch and 1,500 pounds per square inch. Relatively pure hydrogen is, of course, particularly suitable but in the plant operation of our process hydrogen containing impuritiessuch asl methane is often available at much lower cost and can beused effectively as long as the hydrogen content of the gas is above mol percent, in which case the hydrogenpressureV previously mentioned would be the hydrogen partial pressure rather than thetotal gas pressure.

The reactants are preheated to temperatures between 200 F. and"400 F. in heaterv 25 and are directed via line I8' and lines 9| and I9 to d reactor C which is charged with newly regenerlyst can be prepared in reactor 21 by' contacting therein a parafnic naphtha from line 28 and anhydrous aluminum chloride from hopper 29erl other source. Hydrogen chloride or other activator can be added from a source (not shown) via line 30 to the complex generator. The reactants are preferably maintained at a temperature' within the range of about 50 F. to about 225 F., preferably about 100 F. to about 150 F. and the generator is s'o arranged that reflux will be provided in the upper part of the reactor. 4

external to the system. The unconverted hydrocarbon and light gases formed can be withdrawn from catalyst generator 2l through ,line 35. A

` particularly suitable complex andone which we prefer is made bythe reaction of commercial isooctane containing a large amount of 2,2,4- trimethylpentane with ,aluminum chloride using hydrogen chloride as an activator. ,The fresh complex is withdrawn from generatori" through valve 3B and line 31 joining line 26.

i The fresh or regenerated catalyst is Preferably added to reactor C via line 38 andheater -39 (which may be a heat exchanger) and valved line 40. As was previously pointed out, the fresh catalyst apparently requires an induction period or period of contact with a feed stream before optimum? activity is obtained and this induction period can be materiallyshortened if carried out at elevated temperatures. Accordingly. the feed stream from line I8 and lin I9 is contacted with thev catalyst in reactor C at a temperature of about 300 F. to about 400 F.,` preferably about 350 F. for af period of about two hours. The fresh feed bubbling through the tower passes' overhead through line 4I and is directed via valve I2 and lines 43', Il and 45 to line 66 leading to reactor A. Reactor A contains catalyst ywhich is no longer suitable for catalyzing the isomerization reactionunder optimum conditions,

i. e., at temperatures within the range of from 200 to 250 F. wherein a high equilibrium octane `number can be obtained. Reactor A is also maintained at a temperature within the range of about 300 to 400 F., preferably about 330 F., and under the previously described conditions of pressure, activator concentration, etc.

The 'products from the top of reactor A are I .withdrawn through line 56 and directed to reactor B through lines 5B, 59, 62, 64 and 65, valve 60 in line 58 being closed and valves 5l, 6i and S3 in lines 58, 59 and Sl respectively being open. Reactor B is maintained within an optimum temperature range of from about 200 F. to about 250 F., preferably about 212 F. and Aunder the previvously described conditions of pressure, activator concentration, etc., and therein a high octane number is obtained, or inother words the iinal d conversion of the parailinic hydrocarbons in the and valvell in line $3, anddirecting the feedstock directly to reactor A by opening. valve 82 in line 83 which joins line ttf Reactor C is meanwhile lowered in temperature and as soon as the catalyst in reactor B is no longer e'ective for obtaining the maximum amount of conversion of feed stock to a high octane number at that temperature, the temperature in reactor B is gradually raised until it is at the temperature previously used in reactor A. The feed stock to reactor A is then cut oil` by closing valve 82 in line 83 and the fresh feed directed to reactor B by opening valve 8 in line IB and valve 94% in line 85 which joins line t5, valve 86 in line I8 being closed. Reactor C is now put on stream as the optimum reactor, the reaction therein being carried out at temperatures within thev range from 200 to 250 F., the feed stock passing from line 2l through valve 39 from line lt to line it by opening valve 89 therein and thence through valve 99 and line 9| to' line i9.

Meanwhile, the catalyst has been withdrawn y from reactor A by opening valve @l in line 98 and can be discarded by opening valve 99 in line lil@ or used for the production of hydrogen chloride by treatment with water, steam, or-sulfuric acid. If water is employed it should be used in less than stoichiometric amounts in order that the recovered hydrogen chloride will be substanvbe added to reactor A from line 2S via valve di and lines tt, t@ and 9@ passing through heater 5|, if desired, so that itis heated to the proper temperature. Catalyst canalso be introduced concurrently with the incoming feed stock by introducing it through line t2 in which case, valve 56a in line 50 will be closed and valve 52a in line 52 will beopen; Another alternative isl to-add the catalyst directly to the incoming charging stock via line '5A and valve 55 leading from line 2t to line AG. It is now necessary that the catalyst therein be subjected to the initial induction period. Accordingly, the fresh feed stock in line It ows through valve 92, line 93 and line t6 to reactor A, thence through valve 5i and line 58 and valve til to line i8, and then through valve 8d, and lines 85 and 55 to reactor B which is being maintained at the higher temperatures; then through line 16, valve 88, line 81 and valve 89 to line I8, valve 99, line 9| and line I9 to reactor C (valves 8l and 86 being closed). The products are taken overhead from line di through valve lill and line |02 which joins line 'H9 leading to separator 80.

After the catalyst in reactor B becomes completely spent, valve 8i in line 85 is closed thereby shutting o' the flow of charging stock to reactor B. The charging stock thereupon ows llirst to reactor C as previously described and thence to reactor A, the ultimate product being taken overhead yfrom the reactor A by opening valve |03 in line It@ which joins line 19 leading to separator 80, reactor C being at the higher temperature and reactor A at the optimum temperature.

The spent catalyst from reactor B is withdrawn by opening valve IW inline it@ and can also be discarded or sent for hydrogen chloride recovery by opening valve |01 inline w8. Fresh catalyst can be added from line 26 via iine t6 and valve 51 passing through line G8, heater 69 and valved line lll. Again, if it is desirable to pass the stock concurrently with the catalyst, the catalyst can be added through line lll, valve 12 in line 10' being closed and valve 'i3 in line li being open. Alternatively, the catalyst can also be added directly to the product stream entering through line 65 by opening valve "lli in line 'l5 which leads directly from line 26 to linet. Reactor B is nowput in the line first, prior to the high temperature reactor. To accomplish this, feed stock is directed via line I8, line 35, valve 8e and line S5 to reactor B, the hydrocarbon and activators then passing through line it, valve 88, line 8l, line |09, valve il@ line L| il, and line @Il to line i5 and line t6, thence through.V reactor A, line '56, valve 5'?, line 58, line 59, valve ti, line t2, line H2, valve H3, and line Hd to line i9 leading to reactor C.

It will thus be seen that the system is so arranged that any reactor can be cut out of the system for the installation of fresh catalyst, placed back on stream for an induction period at any point in the system, and that the products can'ow from the reactor at the high temperatureoto the reactor of optimum temperature and thence to separator 80. By the proper manipulation of the valves the various products are directed via the various lines to and from each of the reactors.- Catalyst can also be withdrawn from reactor C by opening valve H5 in line H6 and discarded or used for hydrogen chloride production by opening valve in lineA lit. Fresh catalyst can be added to reactor C either through line lit' or, if to be used concurrently, through line 92, valve 93 in line lil being closed and valve 9d in line 92 being open. Alternatively, it can also be added directly to the charging stream by opening valve 95 in line 96 which joins line i9.

It is also perfectly possible to recycle the catalyst on each of the reactors or to recycle a portion of the catalyst and withdraw the remainder for regeneration or discard. In order to obtain recycle on reactor A, line H9 having valve |2 therein leads from line @d to line t9; in reactor B recycle is obtained by opening valve llil in line |22 which leads from line |96 to line 68; and in reactor 'C recycle is obtained by opening valve |23 in line IM which leads from line H6 to line 39. By the proper adjustment of valves 99,425 and |26 in lines 98, it@ and H6 and valves |20, |2I and |23 in lines H9, |22 and EN, the proper proportion of recycle and withdrawn catalyst can be obtained.

The isomerization products in line 'I9 are passed to separator 80 which can be a Warm settler drum `or preferably a cool settler at reduced pressures,

wherein dissolved as well as mechanically entrained catalyst is removed and hydrogen chloride and other gases discharged for recycle. Actually, although separator 89 has lbeen shown as a single vessel, it; may be a series of vessels, prei'- erably arranged at decreasing temperatures and pressures, with means for the recovery of catalyst from the base of each and. the discharge of gaseous hydrocarbons from the top of one or more of them. Preferably, if the cool settler is used. a temperature of F. or less and a pressure within the approximate range of 100 to 300, for example about 250 pounds per square inch, is employed. 'I'he catalyticmaterial withdrawn from the base of separator 80 through line |21 can, of course, |be discarded by opening valve |28 in line |29 but preferably is returned for the formation of additional amounts of complex by opening ,valve |30 in line I3| leading to catalyst generator 21. If for example a warm settler and a cold settler are each employed it is quite possible that a considerable amount of complex will be obtained in the warm settler and this can be recycled directly to the reaction zones by opening valve |32 inline |33 which joins line 26, valve |34 in line |3I thereon being closed.

The released gases leave the top of separator 30 through line |35 and can be vented .by opening valve |36 in line |31. Preferably,lhowever, they are recycled via line |38 and line I6 to line I4 leading to absorber I3 so that the hydrogen chloride present with or without the hydrogen can rbe reused in the reaction. Alternatively, of course, it is possible to recycle the hydrogen plus hydrogen chloride to the hydrogen inlet line 22 by opening valve |39 in line |40 which joins line 22, valve I4| in line I6 being closed, but/generally this is not too desirable, since the amount of hydrogen added can be regulated to a slight excess above that normally required to maintain the activity and catalyst life of the complex to the greatest possible extent while the hydrogen chloride is seldom consumed to anywhere near the same extent, and therefore it is much more advantageous to have the hydrogen chloride reabsorbed with the elimination of unwanted light gases which may build up in the system than to attempt to recover the hydrogen chloride and hydrogen together.v

The products in separator 80 are withdrawn through liney |42 and are desirably sent to a hydrogen chloride stripping tower |43 provided with heating means at its base, and may have a reflux arrangement (not shown) for the recovery of butane and heavienhydrocarbons. This stripping column can be operated at a pressure of about 200 pounds per square inch with a top temperature within the approximate range of 100 to 150 F. and a bottom temperature within the approximate range of 3D0-to 400 F. The removed hydrogen chloride togetherwith released gases such as hydrogen, methane, ethane, etc., is taken overhead through line- |44 to line I6 which will return the gases preferably to the aibsorber system for the recovery of hydrogen chloride, or if necessary. they can be 'discarded through line |44a.

The liquid products are withdrawn from the .base of hydrogen chloride stripper |43 by line .|46 and directed to a wash tower |46 wherein any hnal traces of acid gas or catalyst are removed.

, The wash tower may be provided with suitable homes, trays or bubble plates for enecting intimate contact of the upiiowing products with .the caustic solution introduced through line |41 and withdrawn through line |46. The washed products pass overhead through line |49 and can be withdrawn as such byV opening valve |50A in line ISI but preferably are subjected to fractionation to separate a desired isoma "drom- .heavier product will be withdrawn via valved line |55 while the desired light product is obtained from line |56. Lighter gases which might still be present Ipass overhead through line |51 and can be vented through valved line |58. Alternately, all of the ir'somate can be withdrawn either through line |55eor line |56, and the'light gases taken overhead.

Ammugh we have thus far discussed a process in which catalyst is maintained within a reactor until substantially exhausted for promoting the isomerization of the feed stock further, the same effect can :be obtained in another manner; for example, it is possible to use reactor C always as the induction period reactor, reactor A as the high temperature reactor, and reactor B as the optimum condition reactor, and to direct the catalyst from one reactor to the other in that sequence either continuously or batchwise. In this event fresh feed stock will always be directed to reactor C for the induction period via lines I8, 9| and |9 and the hydrocarbon plus activator and hydrogen taken overhead through line 4| and directed to reactor A via lines 43, 44, and 46. Tha products from reactor A pass overhead through lines 56, 58, 59, 62, 64 and 65 to reactor B and the products are withdrawn therefrom via lines 16, 11 and 19 to separator 60. Meanwhile, fresh catalyst is added izo-reactor C from line 26 either via lines 38, 39 and 92 or via line 96, the conditioned catalyst from reactor C lbeing directed via lines H6, |59 and |60 to lines 66 and 1I or 10 to reactor B which is maintained at the optimum temperature. The catalyst `which is no lon-ger effective at the lower temperatures for obtaining optimum conversion in reactor Bis withdrawn via line |06 and directed to reactor A via lines |6I, |62, 49 and 5| to reactor A wherein it is maintained at the higher temperature for the conversion of paramnic hydrocarbons to isopar- `4l) afiinic hydrocarbons in the feedstock. Other sequences can, of course, fbe. arranged, lines having been indicated for that-purpose, so that any or the reactors can be employed at any of the temperatures and with any of the fed stocks previously described. Recycle lines are also provided for reactors A and B as well as C, although, gener- .ally speaking, since reactor C is only the prelim- -a definite temperature for a denite purpose as compared with the process wherein the temperature of the reactor is changed as conditions warrant. For example, it has beenfound that when operating at the higher temperature. much less reactor volume is required. Every given feed stock basan equilibrium octane number which is the highest constant octane numberobtainable 'on that feed stock at a specic temperature,

andis not affected by increased time of contact or increased ratio of catalyst-to-oil or any other s, similar variable. For example, the equilibrium octane number ofa light naphtha feed stock having a point 'of about 150 F. will be 85 at 212 F. and 82 at 330 F.; the feed having had an initial octane number of 70. Accordingly. from this it will be seen that no amount of treatment in reactr` A at 330 F. will produce a product having an octane number greater than 82.

u 2,375,321 .On the other hand, .operating at 212 F., where'it is possible to obtain the higher octane number, requiresv a much greater time of contact. Accordingly, if reactor n' is maintained at a temperature of about 330 F. and reactor B at a temperature of about 212 F. the volume of reactor A can be substantially smaller than that of reactor B. Thus by carrying out the reaction lin two stages' at the temperatures described, with a proper relationship between the size of the stages, it is possible to decrease the total reaction space from that which would be necessary for a single stage reactor operating at a temperature sumciently low so that the desired high octane num- -ber of 85 might be attained. Reactor C can also be of' a smaller volume than reactor B; since reactor C -is used primarily for conditioning the catalyst for operation in reactor B and since the induction period of the catalyst is substantially less than that of the life ofthe catalyst inv reactor B, it will be necessary to maintain reactor C only so large-as is necessary to afford a con- `\tinuous supply of'optimum condition catalyst to A reactor B. The feed stock to reactor C' can be cut in and out as necessary, or the feed stock to reactor C will be a minor portion of the total feed stock to reactor B using only so much of the fresh feed stock as is necessary to condition the fresh catalyst.

Although we have shown two reactors plus av conditioning reactor, it should be notedv that our invention is not so limited and that more stages can be employed. The additional reaction zones' may also be of diiferent sizes, since the addition of further reactors operating at an intermediate temperature would also conserve reactor volume.

The reactors may advantageously take the form of towers with a deep pool of catalyst thereih, although, as we have stated before, a coil or tubular reactor can be substituted therefor. Stirring devices may be 'employed to insure the intimate contact between reactants andthe catalyst. In addition, bailes may be supplied to any or all of the tower reactors to insure a longer time of contact within the tower, or the tower can be divided into annular "spaces by concentric bales which cause the naphtha to flow over and under the cailles and thus travel in a tortuous as by contact with isooctane or similar hydrocarbons in the substantial absence of hydrogen and A at a lower'temperature than that at which thex original reaction was carried out, the usefulness of the catalyst is again increased, and greater yield of product per unit of catalyst results. Other meansof catalyst regeneration can also be employed.

Although we have described one preferred em. bodiment of our invention, it should bepointed out that this is by way of illustration rather than by way of limitation and that we intend to be l. A process of isomerizing a paraihnic hydrocarbon feed of the naphtha boiling range whichy processcomprises contacting a portion of the vhydrocarbon feed with a fresh liquid aluminum chloride-hydrocarbon complex somerization catalyst in a preliminary step at a temperature within the approximate range of 300 F. and 400 F. and in the presence of an effective amount of an hydrogen Vchloride promoter for a period of time approximating the induction period of the cata-` reaction step at a temperature lower than the temperature of said rst reactionvv step and within the approximate range of u201)" F. and 300 0., the catalyst from said second reaction step being used in said rst reaction step after the activity path through the liquid catalyst thus increasing over the bubble plate and thus increase the time of contact between the hydrocarbons and the cat alyst. Additional variations of the design of thev tower can ne employed without departing from the spirit oi this invention.

instead of withdrawing the spent catalyst, as previously described, either for discard or the recovery of hydrogen chloride therefrom, itis quite 'within the contemplation of our invention to regenerate the catalyst as. by treatment with hydrogen in the substantial absence of added naphthas. Also, it has been found that very often a catalyst which has become too viscous for use in the reactor, due either to the removal of hydrocarbons joined to the aluminum chloride in the original complex or by the addition of addi tional hydrocarbons to the complex, still retains a desirable amount of catalytic activity which cannot be utilized because of failure properly to thereof at the lower temperatures has decreased.A

2. A. process of isomerizing a yparalnic hydrocarbon feed of th naphtha boiling range which process comprises contacting a portion of the hydrocarbon 'feed with a fresh liquid aluminum chloride-hydrocarbon complex isomerization catalyst in a preliminary step at a temperature of about 300 Fand in the presence of an eective amount of an hydrogen chloride promoter for approximately two hours so as to form a catalyst of high activity, then contacting the hydrocarbon feed with a partially spent catalyst complex in a first reaction step at a temperature of about 330 F. and finally contacting the hydrocarbon products from said ilrst reaction step with the catalyst of high activity in a second reaction step at a temperature of about 212 F., the catalyst from said second reaction step being `used in said first reaction step after the activitythereof at the lower temperature has decreased.

3. A process as claimed in claim l wherein the paramnic hydrocarbon feed of the naphtha boiling range consists predominantly of C4 to Cs parailnic hydrocarbons.

4. A process of isomerizing a parailinic hydrocarbon feed of the naphtha boiling range, which process comprises contacting a portion of' the hydrocarbon feed with a fresh liquid aluminum chloride-hydrocarbon complex isomerization catalyst in apreliminary step at a temperature of about 300 and in the presence of an eHec- Y,

riod of the catalyst to form a` catalyst of high activity, then contacting the hydrocarbon feed Wishing" t0 with a'partially spent catalyst complex in a lfst the temperature of said preliminary step, the reaction step at a temperature below 400 F. but catalyst from said second reaction step being used rmt' IOWer than the iempeatllle 0f Said Prelimiin said first reaction step after the activity therenary step, and finally Contacting the hydrOCalof at the lower temperature has decreased;

bon products from said rst reactionjstep with 5 the catalyst of high activity in a. second reaction CECIL W NYSEWANDER.. step at an isomerization temperature lower than NATHAN FRAGEN. 

